GB2091756A

GB2091756A - Production of acetylene

Info

Publication number: GB2091756A
Application number: GB8138433A
Authority: GB
Original assignee: GAF Corp
Current assignee: GAF Corp
Priority date: 1980-12-23
Filing date: 1981-12-21
Publication date: 1982-08-04
Also published as: GB2132183A; US4367363A; JPS57131730A; GB2091756B; DE3150340A1; AU7775381A; GB8400329D0; AU5647786A; GB2132183B

Description

1 GB 2 091 756 A 1

SPECIFICATION

Production of acetylene This invention relates to the production of acetylene from a feed stream which contains impurities such as HM, H2S, CS,, CO and CO,.

The acetylene rich output from reactors such as plasma art reactors, radiant energy reactors and the like, have great value if the acetylene can be separated as a high yield product, free of impurities and if the process can be run under reasonable economic conditions.

The removal of acidic gaseous mixtures is dis- closed in U.S. Patent 3,565,573, the removal Of C02 and the washing of acetylene with an organic solvent is disclosed in U.S. Patent 3,686,344, acetylene separation is disclosed in 3,325,972, the removal of hydrogen cyanide from acidic gases is disclosed in U.S. Patent 3,935,188 and the removal of H2S and C02 is disclosed in 3, 829,521, as examples of available technology. However, an integrated system is required to achieve the desired end result within the frame work of the economic restraints of a commer- cial system.

In accordance with the present invention, pure acetylene is produced from a feed stock which includes such impurities as H2S, SC2, CO and C02. In a first stage acid gases such as HM and H2S are selectively removed using an amine or an alkyl pyrrolidone solvent and C02 is removed using a caustic scrubber. In a second stage, the same solvent is used to remove CS, and produce an acetylene rich product. In a third stage, the remaining gases are sub- jected to hydrogenation and desulferization prior to methanation and then recycled to a coal to acetylene reactor.

An embodiment of the invention will be described and exemplified with reference to the drawings, wherein:

Figure 1 is a schematic representation of a first part of the acid gas clean-up portion of the process; Figure 2 is a schematic representation of a first part of the sweet gas clean-up portion of the pro- cess; Figure 3 is a schematic representation of a second part of the sweet gas clean-up operation; Figure 4 is a schematic illustration of the hydrogenation, desulfurization and methanation portion of the process; Figure 5 is a schematic illustration of the organic solvent clean-up portion of the process; and Figure 6 is a schematic illustration of the overall process.

Referring to the drawings, the acetylene gas separation process includes four stages of operation. The gas first goes through an acid gas removal section where HCN and H2S are selectively removed utilizing 5% NIVIP in water as the absorption solvent (The initials NMP denote N-methyl pyrollidone). The CO, is then removed with a caustic scrubber. The gas then passes to the second stage.

In the second stage, CS2 is removed using NMP as the solvent, and acetylene is separated from the other gasses (CH2, C2C2, C2H4, C3H,, C41-12, CO).

In the third stage, the remaining gasses are hydrogenated, desulfurized and methanated. The treated gasses are then recycled to the source. Any hydrogen blowdown would occur at this point.

The fourth stage is the NIVIP cleanup, where NIVIP is purified via limestone treat and distillation.

The feed 10 can be the output from an acetylene gas generation system such as an arc generator of the type disclosed in U.S. patents such as 3, 333,027, 4,105,888 and 4,010,090 and high temperature reactors of the type disclosed in U.S. patents 3,933,434 or other acetylene generator systems, such as 3,384,467. Similarly, otherforms of acetylene generators can be used as the source of the feed stream.

The feed stream 10 is treated in an acid absorber 102 with an organic solvent, which is preferably an alkylamine or alkyl pyrrolidone such as Nmethyl pyrrolidone.

The organic solvent and almost all of the HM and H2S are removed as bottoms, via line 12 and feed to an acid stripper 101. The HCN from the top of the acid stripper 101 is condensed in a heat exchanger 105 and withdrawn from the system via stream 13. In the divider 104 HS and organic solvent are separated a n d H2S i n th e strea m 14 is with d rawn f ro m th e system. The organic solvent, along with any additional makeup solvent, as required, is returned to the acid absorber 102, via the line 15.

The stream 16 contains the overheads from the acid absorber 102 and is rich in the H2, CO, CH4, C2H4, C21-12, C31-14, C41-12, CS2, and C02 of the feed stream. The materials of the stream 16 are treated in the caustic chamber 103 with a caustic such as N,,OH to separate the residual HM and most of the residual H2S and C02 from the system via line 17. The overhead effluent from the caustic scrubber 103 removed via line 18 is compressed in a compressor 106 and fed to a CS2 absorber 107 in which an organic solvent is used, which is preferably, the same solvent as used in acid absorber 102. The CS2 and C.,H2 is essentially removed, along with most of the C3K, and some C2H., via line 20 and delivered to CS2 stripper 110 after cooling in a heat exchanger 109.

Stream 30 from the bottom of the CS2 stripper containing C3H4 and Nmethyl pyrrolidone is cooled in the heat exchanger 116 and treated in the final stripper 131.

The top stream 23Tfrom the CS2 stripper 110 is heated in a heat exchanger 111 and split into a portion 23R which is returned to the stripper 110 as reflux, and a portion 23 which is compressed in the compressor 112 cooled in the heat exchanger 113, passed through a flash tank 114 where the CS2 along with minor amounts of C21-12, C3H4 and C4H2 are separated as stream 26, from the more volative constituents including H2, CO, CH4, C2H4 and C2H2 in stream 27 which is fed to mixer 115 along with streams 28 and 32 to produce a combined stream which is high in H2, CO, CH,, C2H4 and C2H2.

The main absorber 120 uses an organic solvent, preferably the same solvent as used in other parts of the system, to strip most of the C2H2 from the com bined streams 27, 28 and 32.

The bottoms 35 from the main absorber 120 are 2 GB 2 091 756 A 2 flash separated at 122 into a gas stream 33 and a liquid stream 34.

The gas stream is compressed in compressor 121 and recycled as stream 32. The liquid stream 34 is heated in the heat exchanger 123 and fed to the main 70 stripper 124 from which acetylene product is removed as an overhead product.

The bottoms 47, containing the organic solvent and primar,.y C2H4 with some C3R,, CJ-12 and CS, is heated and flashed to separate a liquid stream 50 containing most of the organic solvent and C2H4, C3H4 and minor amounts of CH2, CS2 and organic solvent. The gaseous stream 49 is cooled in heat exchanger 130 and returned to the bottom of the 15.-nain stripper 124.

The bottoms 50 are fed to the final stripper 131 along with the heated bottoms 31 from CS2 stripper 110, and a heated portion of the overhead from the main absorber 120.

The bottoms 51 from the final stripper contains most of the organic solvent and CH2, along with minor amounts of H2, CO, CH4, C2H4, C21-12, and CS2. The overheads 62 is heated in a reflux heat exchanger 138.

The compressed overheads 62 from the final stripper 131 and a portion of the overheads 42 from the main absorber 120 are heated in a heat exchanger 140 and treated in a catalytic hydrogenator 141, thereby converting a portion of the hydrogen and the C2H,, C3H4 and C4H, into C2H6, C3H6 and C4H1c). The hydrogenated stream 65 is passed in heat exchange with the stream 69 from the desulferizer 145 and then a portion of the stream 72 from the methanators 150 and 151 priorto being processed in the catalytic desulferizers 144 and 145.

The desulferized stream 70 is fed to the pair of methanators 150 and 151 where hydrogen and CO are converted to CH,, and water. A first portion 78 of the methanator output stream 72 is recycled through a heat exchanger 146 where it heats a recycle stream 77 which is fed to the heat exchanger 143 and then combined with another portion 73, of the stream 72 and fed to an economizer boiler 149 where it gives up heat and then passes to a dehydrator 148 where water and residual organic solvent are removed via stream 76.

Looking once again to the bottoms 51 from the final stripper 131, a portion 52 is subjected to a limestone treatment in the divider 132 and subjected to an organic solvent clean up in a distillation column 137 and another portion 56 is mixed with the stream 55 from column 137. The mixing occurs in a filteration mechanism 133 and the predominant organic solvent resultant stream 57 is heated at heat exchanger, compressed at the compressor 135 and divided into two streams. A first stream 60 is fed to the main absorber 120 and a second stream 61 is fed to the CS2 absorber 107.

The bottoms 54 foam distillation column 137 is essentially organic solvent and a low molecular weight polymer while the overheads 53 is a mixture of hydrogen, C,,H2 and organic solvent. EXAMPLE The following example sets forth the process parameters and the material and energy balances for the separations of a plurality of end products and impurities from the output of a plasma arc reactor of the type disclosed in U.S. patents 4010090 and 4158637. The Tables 1 through VIII set forth the compositions of each stream in Lb. Mole per hour, unless otherwise specified, temperatures, pressures, and in some instances, enthalpy and liquid or vapor flow rates. 75 The reflux heat exchanges 111, 125,105 and 138 specifications are set forth in Table IX. The scrubber, absorber and stripper column specifications are set forth in Table X. The Wmethyl pyrrolidone distillation tower 137, is 80 a 10 tray tower operating at 400OMM Hg., at 360'F and handles 100 lb. moles per hour of feed. The hydrogenator, desulferizers and methanators employ catalysts which are commonly used forthe respective function. It should be noted thatthe pro- cess design calls for conventional equipment as well known in the art.

TABLES

The foregoing described process is noted to depart from the state of the art in a variety of ways.

For example, the system generates acetylene from gassified coal with acid gas levels in the 1 to 5% range and acetylene in the 3 to 20% range and the sweet gas system uses pressure vessels only, and thus avoids the expense associated with vacuum systems, H2S, CS2, HCN and C2H2 are recovered as pure, unreacted products. Furthermore, the unique equipment train and process sequence uses an alkyl amine or an alkyl pyrrolidone in dilute form to remove H2S and HCN as well as in the sweet gas section and thus is a one solvent system.

As a further feature, overhead gases are used to reduce concentration gradients in the stripping operations and thus enhance stripping. The use of conditioning equipment in the recycle stream boosts the yield by enabling the recycling of the hydrocarbon stream. The clesulferization priorto methanation avoids the poisoning of the methanator catalyst.

Z -1 1 3 GB 2 091 756 A 3 TABLEI ACID GAS LB. MOLES1HR.

Stream 10 12 14 13 15 16 18 17 65 H2 12,225 4.28 4.28 0 0 12,221 12,221 0 0 CO 620 1.43.57 0.15 620 620.29 25.7 CH4 171.29.29 0 0 171 171 0 0 C211, 128.3 0 0 0 0 128.3 128.3 0 0 C211, 1,677 17.1 17.1 0 0 1,660 1,660 0.6 51.3 C3H4 11.4 0 0 0 0 11.4 11.4 0 0 C4H, 11.4 0 0 0 0 11.4 11.4 0 0 CS, 28.5 0.15 0.15 0 0 28.5 28.5 0.15 12.8 5% NMP in H,0 0 9,975 0 0 9,975 0 0 0 0 10% NaOH in H20 0 0 0 0 0 0 0 114 10,260 CO, 12.8 0.29 0.29 0 0 12.8 1.43 11.4 1,026 H2S 28.5 27.1 27.1 0 0 1.43 0.29 1.43 128.3 HM 178 178 0.29 178 0.29 0.6 0 0.6 51.3 Ternp,'F 77 77 95 120 90 77 77 89 77 P, psia 29 29 14.7 14.7 29 29 29 14.7 14.7 TABLE11 SWEET GAS LB. MOLES1HR.

Stream 19 20 21 22 28 27 26 25 24 23 H, CO CH4 C2H4 C,H2 C^ C4H2 CS2 NMP Ternp,'F P_ psia Enthalpy, BTU/Hr. Vol. Flow Rate (gpm) liquid Vol. Flow Rate (ft.3/sec.) vapor 12,220 619 171 129 1,660 10.9 11.4 28.2 5.9 0.34 0.23 1.0 233 7.6 12.0 28.2 1138.

79.9 225 5.9 0.34 0.23 1.0 233 7.6 12.0 28.2 1138.

72.3 14.7 5,9 12,215 0.34 618 0.23 171 1.0 128 233 1,427 7.6 3.3 12.0 0.1 28.2 0.1 1138. 0.7 266 86.5 14.7 225 77 225.49 X 107 -26 x 108 -26 x 108 -.15 x 108.58 x 107 250 236 237 106 255 12.9 3.8 3.4 181 3.1 0.2 1.5 13.6 41.3 106 2.80 0.04 255 0.02 12.9 0 3.8 0.07 3.5 51.5 233 7.6 7.6 4.8 5.0 25.7 27.3 0.7 0.7 10 10 225 225 225 -.13 X 106 -33x106 -.87 X 106.35 X 107.33 X 106 9.4 9.4 255 12.9 3.8 3.5 233 7.6 5.0 27.3 0.7 677 225 255 12.9 3.8 3.5 233 7.6 5.0 27.3 0.7 100 14.7 2.80 8.28 62.1 4 GB 2 091756 A 4 TABLE111 SWEETGAS LB. MOLES1HR.

39 38 37 42 41 43 29 -11 Stree m, 30 31 H2 CO CH4 C2b4 C2H2 C3H4 C4H2 Cs 2 1WP 1137 Temp,'F 242.6 P, psia 14.7 Enthalpy, -.17 X 101 BTU/Hr.

Vol. Flow Rate (gpm) liquid Vol. Flow Rate (ft.3/sec.) vapor 0.3 0.3 250 0.01.01 12.7 0.01.01 3.5 0.02.02 2.5 0 0.03 0 0 0 7.0 7.0 0 8 0 1137 0 350 266 14.7 14.7 _1.09 X 107.44 X 106 250 12.7 3.5 2.5.03 0 0 0 0 77.8 14.7.86 x 105 12,549 635 180. 143. 2,249 6.8.34 1.7 1.25 112.4 225.86 X 105.39 X 107.30 X 10'.30 X 106.21 X 107.93 x 107 250 11,357 12.7 575 3.5 159.

2.5 116.

03 1.7 0 0 0.1 0.02 0.50 77.2 77.2 225 225 874 44.2 12.2 8.9.1 0 0 0 0 77.2 225 874 44.2 12.2 8.9.1 0 0 0 0 77.9 14.7 874 44.2 12.2 8.9.1 0 0 0 0 350 14.7 8 z 237 250 39.4 29.3 1.93 87.6 6.7 102 154 118 TABLEIV SWEETGAS LB. MOLES1HR.

Stream 36 44 35 46 48 32 33 34 45 47 H2 12,471.63 79.6 0 0 79.0 79.0.57.55 0 CO 632.03 4.0 0 0 4.0 4.0 0 0 0 CH4 174.09 3.8 0 0 3.6 3.6.11.11 0 C2H4 2.1 154 2,247 191.0 126 641. 641. 1,606 1,606 191 C2H2 129. 1.60 14.4 0 0 12.8 12.8 1.6 1.6 0 C3H4.0 5.1 6.8 5.1 4.0.57.57 6.3 6.3 5.1 C4H2.1 1.17 12.0 11.7 1.0.11.11 11.9 11.9 11.7 CS2.03 0.88 25.1 1.7.2.11.11 2.4 2.4 1.7 NMP.57 1.86 16,940 17,074 137..68.68 16,939 16,939 17,074 Temp,'F 77.3 100 90 336.3 300 572.7 87.8 87.8 300 350 P, psia 225 14.7 225 14.7 14.7 225 14.7 14.7 14.7 14.7 Enthalpy.,.42 X 107 1.05 x 106 -38 x 109 -.21 X 109.51 x 106 -.49 x 107 -. 43 X 107 -38 x 1011 -Ag X 1011 -.16 x 109 BTU/Hr. Vol. Flow Rate (gpm) liquid 3502 3659 21.6 3431 3557 3720 Vol. Flow Rate (ft.3 /sec.) vapor 84.4 159 25.4 10.8 82.0 256 43.4 GB 2 091 756 A 5 TABLE V SWEET GAS LB. MOLE1HR.

Stream 49 50 62 51 52 56 55 60 57 61 H2 co CH4 C2H4 C2H2 C3H4 C4H2 CS2 NMP Temp.,'F P., psia Enthalpy, BTU/Hr. Vol. Flow Rate (gpm) liquid Vol. Flow Rate (ft.3/sec.) vapor 0 0 0 0 0 0 0 0 126 65 4.0 3.0 1.0 9.6 2 1.5 137. 16,937.

350 350 14.7 14.7 19 X 107 -.16 X 109 3722 43.6 870 44.6 12.2 8.7.2 64.2.01 3.0 0 4.0 12.5 1.5.8 7.2 18,070.

324.8 14.7 14.7 39.2.1 02.39 0.2 0.1 0.2 0.01 0 0 08 12.4 0.8 114 17,960.

324.8 324.8 14.7 14.7 0 0 0 0 0 0 0 0 114 350 14.7 3.7.39 23.2 11.1 23.2 01.01 0 0 11.6 12.4 80.8 16,937 18,070.

77.5 325 225 14.7 2.01 0.01 0 0.79.06 1,139. 77.5 225.89 x 10r, -19 x 109.13 x 107 -.19 x 109 -1.09 x 106 -38 x 109 -.19 x 1011 -.25 x 1011 3722 24.6 3882 TABLE W SWEET GAS LB. MOLES1HR.

Stream 59 58 53 54 H2.39.39.02 0 co.2.2 0 0 CH4.1.1 0 0 C2H4.2.2 0 0 C2H2.01.01 0 0 C3H4 0 0 0 0 C4H2 12.4 12.4.08 0 CS2.8.8 0 0 NMP 18,070 18,070.02 2.3 Polymer (NW 200) 0 0 0 2.3 Temp.,'F 77.5 77 90 424 P., psia 225 14.7 14.7 14.7 Enthalpy, BTU/Hr. -Al X 109 -Al X 19 Vol.FlowRate,gpm 3,476 3,486 25.0 3257 3907 219.

GB 2 091 756 A 6 TA BL E VII SWEET GAS LB. MOLES1HR.

63 64 65 66 67 68 69 70 79 71 72 H, 12,227 12,227 11,948 11,948 11,948 11,948 11,948 11,948 15,455 27,037 25,547 CO 618 618 618 618 618 618 618 613 0 618 0 CH4 171 171 171 171 171 171 171 171 1,210 1,381 2,000 C21-14 125 125 0 0 0 0 0 0 0 0 C2H2 65.9 65.9 0 0 0 0 0 0 0 0 C31-14 3.0 3.0 0 0 0 0 0 0 0 0 C41-12 4.1 4.1 0 0 0 0 0 0 0 0 CS2 (sulfides) 1.5 1.5 1.5 1.5 1.5.6.01.01.01.025.03 H20 x x x x x x x x 947 947 1570 CA x x 190 190 190 190 190 190 292 482 482 C4Hjo x x 4.1 4.1 4.1 4.1 4.1 4.1 6.3 10.4 10.4 C,H, x x 3.0 3.0 3 3 3 3 4.6 7.5 7.5 NMP 7.7 7.7 7.7 7.7 7.7 7.7 7.7 7.7 11.7 19.5 19.5 Temp.,'F 50 400 465 600 650 650 650 515 650 515 800 P., psia 45 44 43 42 41 40 39 38 41 38 37 TABLE VIII SWEET GAS LB. MOLES1HR.

73 78 82 83 74 75 77 80 H2 8,133 15,455 1,960 1,960 10,092 10.092 8,853 8,853 CO 0 0 0 0 0 0 0 0 CH4 636 1,210 153 153 790 790 790 790 CS, (sulfides).01.01.01.01.01.01.01.01 H20 499 947 120 120 619 619 0 0 C2H6 153 292 37 37 190 190 190 190 C41-110 3.3 6.3.8.8 4 4 4 4 C^ 2.4 4.6.6.6 3 3 3 3 NMP 6.2 11.7 1.5 1.5 7.6 7.6 0 0 Ternp,'F 800 800 800 515 748 100 100 485 P, psia 37 42 42 38 37 32 30 28 TABLEIX

Reflux TC Number 0 (BTUlhr.) Ratio 'F ill 2.4 x 10.13 14 x 10.32 105.6 138 22 x 10.81 TABLEX

100 80 150 No.

Dia., Theoretical Temp. Pressure, Number Name ft. Trays OF psia 107 CS2 absorber 5.5 10 350 225 CS2 stripper 2.8 5 300 14.7 Main absorber 8.7 10 77 225 124 Main stripper 3.1 10 266 14.7 131 Final stripper 10 5 77 14.7 102 Acid absorber 10 77 29 101 Acid stripper 10 170 14.7 103 Caustic scrubber 5 80 29 k 1 7 GB 2 091 756 A 7

Claims

1. The process of treating a product stream from coal to acetylene conversion operation by separating acetylene grom gases containing carbon disulfide, C, to C, hydrocarbon, hydrogen, carbon monoxide and carbon dioxide, comprising the steps of:

a) treating the gases with an organic solvent selected from the group consisting of alkyl amines and alkyl pyrrolidone; b) distilling a first portion of said gases and separating H,S and I-ICN containing strams from said first portion; c) subject a second portion of said gases to caustic scrubbing and separating HCN, H,S and CO, from said second portion; d) treating said second portion with said second organic solvent and separating said second portion into a first stream containing CS, QJ-12, C3H4, and C2H4 and a second stream containing H2, CO, CH4, C^ and C21-12; e) treating said first stream with said organic solvent and stripping the major portion of said CS2 from said first stream and combining the remaining portion of said first stream with said second stream to form a combined stream; f) treating the CS2 rich portion of said stripped first stream with said organic solvent and separating said CS2 rich portion of said first stream into an essentially CS2 free third stream and a fourth stream containing primarily CS2; g) treating said fourth stream to separate and refine organic solvent; h) recycling the refined organic solvent to the treating operation of step d); i) treating said combined stream with said organic solveritto produce an acetylene product stream and a sixth stream containing C2H4, C,H4, C4H2 and CS2; j) treating said sixth stream with said organic solvent along with said CS2 rich portion of said strip- 105 ped first stream in step f); k) subjecting said essentially CS2 free third stream to hydrogenation, desulferization methanation and recycling the resultant stream to a coal to an acetylene conversion operation.

2. The process of Claim 1, wherein acid organic solvent is an alkyl amine.

3. The process of Claim 1, wherein said organic solvent is an alkyl pyrrolidone.

4. The process of Claim 3, wherein said alkyl pyrrolidone is N-methyl pyrrolidone.

5. The process of Claim 1, wherein steps d) through k) of the process are carried out at at least ambient pressure.

6. The process of Claim 1, wherein hydrogenation is carried out in the presence of a catalyst.

7. The process of Claim 1, wherein the desulferization is carried out in the presence of a catalyst.

8. The process of Claim 1, wherein the methana- tion is carried out in the presence of a catalyst.

9. The process of Claim 1, wherein said product stream contains from about 1 to about 5% acid gases.

10. The process of Claim 1, wherein said product stream contains from about 3 to 20% acetylene.

11. The process of Claim 1, wherein H2S, CS2, HCN and C21-1, are recovered as essentially pu re, unreacted products.

12. The process of Claim 8, wherein desulferiza- tion is essentially completed prior to methanation.

13. The process of Claim 1, wherein the feed stock is the product output stream from a plasma arc reactor and the feed to said reactor is powdered coal.

14. The process for recovering essentially pure acetylene from the gaseous out-put stream from a coal to acetylene conversion process, comprising the steps of:

a) processing said gaseous out-put stream in an acid gas removal first stage by absorbing I-ICN and H2S in an organic solvent, and scrubbing with a caustic agent to remove C02; b) scrubbing said gaseous out-put stream with said organic solvent in a sweet gas treatment second stage and removing essentially pure acetylene as a product; c) in a third stage, first hydrogenating, then desulferizing and then methanating the remaining gases from said second stage and recycling out- put from said third stage to said coal to acetylene con- version process; and d) in a fourth stage, refining said organic solvent from said second stage and recycling organic sol vent to at least one of said first stage and said sec ond stage.

15. The process of Claim 14 wherein said organic solvent is an alkyl amine.

16. The process of Claim 14 wherein said organic solvent is an alkyl pyrrolidone.

17. The process of Claim 16, wherein said alkyl pyrrolidone is N-methyl pyrrolidone.

18. The process of Claim 14, wherein said second stage is carried out at at least ambient pressure.

19. The process of Claim 14, wherein said coal to acetylene conversion process employs a plasma arc generator and said gaseous out-put stream contains from about 3 to 20% acetylene.

20. The process of Claim 19, wherein said gaseous out-put stream contains from about 1 to about 5% acid gases.

Printed for Her Majesty's Stationery Office by The Tweeddale Press Ltd., Berwick-upon-Tweed, 1982. Published at the Patent Office,25 Southampton Buildings, London, WC2A lAY, from which copies may be obtained.